Production of Olefins

ABSTRACT

A process for converting a hydrocarbon feedstock to provide an effluent containing light olefins, the process comprising passing a hydrocarbon feedstock, the feedstock containing at least one C 1  to C 6  aliphatic hetero compound selected from alcohols, ethers, carbonyl compounds and mixtures thereof and steam in an amount whereby the feedstock contains up to 80 weight % steam, through a reactor containing a crystalline silicate catalyst to produce an effluent including propylene, the crystalline silicate having been subjected to de-alumination by a steaming step and being selected from at least one of an MFI-type crystalline silicate having a silicon/aluminium atomic ratio of from 250 to 500 and an MEL-type crystalline silicate having a silicon/aluminium atomic ratio or from 150 to 800.

The present invention relates to a process for converting an oxygencontaining hydrocarbon feedstock to produce an effluent containing lightolefins, in particular propylene.

There is an increasing demand for light olefins, for example ethyleneand propylene, in the petrochemical industry, in particular for theproduction of polymers, iii particular polyethylene and polypropylene.In particular, propylene has become an increasingly valuable product andaccordingly there has been a need for the conversion of varioushydrocarbon feedstocks to produce propylene.

Increasing amounts of stranded or associated natural gas are being foundthroughout the world. It becomes important to valorize these gasreserves, not only as fuel but if possible as a carbon source forchemicals and liquid transportable fuel. One way of doing this is theconversion of natural gas into synthesis gas and consequently synthesisof methanol that can serve as a primary source of other chemicals orliquid fuels.

It has been known for a number of years to convert low molecular weightmonohydric alcohols such as methanol into light olefins, with theeffluent containing ethylene and propylene. Methanol can readily beproduced from methane present in natural gas, which is in abundantsupply, and is in oversupply in some oil-producing regions of the world.There is therefore a need to produce light olefins such as ethylene andpropylene from feedstocks derived from natural gas.

The conversion of a feed containing C₁ to C₄ monohydric alcohols tooleflnic hydrocarbons including ethylene and propylene has been known atleast since the 1970's. For example U.S. Pat. No. 4,148,835 in the nameof Mobil Oil Corporation discloses a catalytic process for converting afeed containing a C₁-C₄ monohydric alcohol, in particular methanol, bycontact of the alcohol, under conversion conditions, with a catalystcomprising a crystallised alumina silicate zeolite having a crystallitesize at least about 1 micron, a silica to alumina ratio of at leastabout 12 and a constraint index within the approximate range of 1 to 12.In particular, the zeolite comprises ZSM 5. The effluent from themethanol conversion includes ethylene and propylene. The problem of theprocess disclosed in U.S. Pat. No. 4,148,835 is that the propylene yieldis not very high and there is a need to increase the propylene yield ofthe conversion process.

EP-A-0123449, also in the name of Mobil Oil Corporation, discloses aprocess for converting alcohols/ethers, especially methanol, intoolefins over zeolite catalysts. Olefin selectivity is enhanced by usingzeolites of crystal size less than 1 micron and which have been steamedto alpha values of not more than 50, preferably 5 to 35. However,although the mixture of olefins produced contains mostly ethylene,propylene and the butylenes with a small pentenes component, there is nodisclosure of a process which has a high propylene selectivity.

DE-A-2935863, and its equivalent U.S. Pat. No. 4,849,753, also in thename of Mobil Oil Corporation, disclose a process for producing lightolefins by catalytically converting methanol over crystallinealuminosilicate zeolites having high silica to alumina ratios attemperatures of from about 350 to 600° C. and at pressures rangingbetween about 1 and 100 atmospheres.

It is also known in the art to convert methanol to light olefins using asilica-alumina-phosphate catalyst, known as SAPO catalysts. It wasconsidered that such catalysts had a higher selectivity to light olefinsthan the alumino-silicate zeolite catalysts employed in, for example,U.S. Pat. No. 4,148,835. For example, U.S. Pat. No. 4,861,938, U.S. Pat.No. 5,126,308 and EP-A-0558839, all in the name of UOP, each disclose aprocess for the conversion of methanol into light olefins, in particularethylene and propylene, using a silica-alumina-phosphate catalyst, inparticular SAPO 34. These processes suffer from the problem that, inparticular, when used in a fixed reactor, the selectivity to propyleneof the catalyst is poor, and additionally too much ethylene is produced,leading to a relatively low propylene/ethylene molar ratio. This lowersthe propylene purity in a fractionated cut containing C₂ and C₃hydrocarbons. Also, as a result of the production of propane, thepropylene purity in a C₃ cut may be low. Furthermore, the propyleneselectivity tends not to be stable over time. There is therefore a needto provide a conversion process which has a higher propylene selectivitythan these known processes.

It is also known to crack catalytically an olefin-containing feedstockusing a crystalline silicate catalyst, for example from WO-A-99/29802(and its corresponding EP-A-0921176) and from WO-A-99/29805 (and itscorresponding EP-A-0921181).

It is further known to use a crystalline silicate cracking catalyst toproduce light olefins such as ethylene. For example, WO-A-98/56877discloses a process for improving the conversion of a light hydrocarbonfeedstock to light olefins comprising the steps of first contacting thehydrocarbon feedstock with a light olefin producing cracking catalyst,such as a ZSM-5 zeolite, and subsequently thermally cracking theunseparated stream to produce additional ethylene.

EP-A-0882692 discloses a process for the production of lower olefinswith 2-3C atoms which comprises reacting a methanol and/or dimethylethervapour and a reaction mixture containing water vapour in a first reactoron a first form selective catalyst at 280-570 degrees C. and 0.1-1 bar;withdrawing a product mixture containing 2-4C olefin and 5C+ hydrocarbonfrom the first reactor; and cooling. The cooled first product mixture isfed through a separator and a second product mixture containing ethyleneand propylene is withdrawn. A 5C+ stream is obtained, which is vaporisedand mixed with water vapour. A ratio of H2O:hydrocarbons of 0.5-3:1 isused. The mixture containing water vapour is fed at 380-700 degrees C.to a second reactor containing a second form selective catalyst. A thirdproduct mixture is withdrawn from the second reactor which contains 50%olefinic components. This product mixture is cooled and fed to aseparator. The catalyst in the first reactor may be a zeolite asdisclosed in EP-B-0448000, a SAPO catalyst as disclosed in U.S. Pat. No.4,524,235 and EP-A-0142156, or a silicalite catalyst as disclosed inU.S. Pat. No. 4,061,724. The catalyst in the second reactor may be azeolite of the Pentasil-type having a silicon/aluminium atomic ratio offrom 10:1 to 200:1, variants of such catalysts being disclosed inEP-B-0369364, a SAPO catalyst or a silicalite catalyst.

It is an object of the present invention to provide a process forconverting oxygen-containing hydrocarbon feedstocks which has a highyield of lighter olefins, and in particular propylene. It is anotherobject of the invention to provide a process for producing propylenehaving a high propylene yield and purity.

It is a further object of the present invention to provide such aprocess which can produce olefin effluents which are within, at least, achemical grade quality.

It is yet a further object of the present invention to provide a processfor producing olefins having a stable olefinic conversion and a stableproduct distribution over time.

The present invention provides a process for converting a hydrocarbonfeedstock to provide an effluent containing light olefins, the processcomprising passing a hydrocarbon feedstock, the feedstock containing atleast one C₁ to C₆ aliphatic hetero compound selected from alcohols,ethers, carbonyl compounds and mixtures thereof and steam in an amountwhereby the feedstock contains up to 80 weight % steam, through areactor containing a crystalline silicate catalyst to produce aneffluent including propylene, the crystalline silicate having beensubjected to de-alumination by a steaming step and being selected fromat least one of an MFI-type crystalline silicate having asilicon/aluminium atomic ratio of from 250 to 500 and an MEL-typecrystalline silicate having a silicon/aluminum atomic ratio of from 150to 800.

Preferably, the MFI-type crystalline silicate catalyst comprisessilicalite.

Preferably, the at least one C₁ to C₆ aliphatic hetero compound is anoxygen containing compound.

Preferably, the hydrocarbon feedstock contains at least one of methanol,ethanol, dimethyl ether, diethyl ether and mixtures thereof.

Preferably, the hydrocarbon feedstock is passed over the crystallinesilicate at a reactor inlet temperature of from 350 to 650° C., morepreferably from 450 to 550° C.

Preferably, the hydrocarbon feedstock is passed over the crystallinesilicate at a WHSV of from 0.5 to 30 h⁻¹, the WHSV being based on theweight of the at least one C₁ to C₆ aliphatic hetero compound in thefeedstock.

Preferably, the partial pressure of the at least one C₁ to C₆ aliphatichetero compound in the feedstock when passed over the crystallinesilicate is from 20 to 400 kPa.

The present invention further provides the use, in a process forconverting a methanol feedstock in a reactor having a reactor inlettemperature of from 450 to 550° C. into an effluent containingpropylene, of a crystalline silicate catalyst which has beende-aluminated by steaming thereby to have a silicon-aluminium atomicratio of from 250 to 500 for increasing the propylene/ethylene ratio inthe effluent.

The present invention yet further provides the use, in a process forconverting a methanol feedstock in a reactor having a reactor inlettemperature of from 450 to 550° C. into an effluent containingpropylene, of a crystalline silicate catalyst which has beende-aluminated by steaming thereby to have a silicon-aluminium atomicratio of from 250 to 500, for increasing the propylene/propane ratio inthe effluent.

The present invention still further provides the use, in a process forconverting a methanol feedstock in a reactor having a reactor inlettemperature of from 450 to 550° C. into an effluent containingpropylene, of a crystalline silicate catalyst which has beende-aluminated by steaming thereby to have a silicon-aluminium atomicratio of from 250 to 500, for enhancing the stability of the catalystover time.

The present invention can thus provide a process wherein hydrocarbonstreams (products) from refinery and petrochemical plants areselectively converted not only into light olefins, but particularly intopropylene.

The hydrocarbon feedstock may be fed either undiluted or diluted withsteam and/or an inert gas such as nitrogen. In the latter case, theabsolute pressure of the feedstock constitutes the partial pressure ofthe hydrocarbon feedstock in the steam and/or the inert gas.

The various aspects of embodiments of the present invention will now bedescribed in greater detail, by way of example only, with reference tothe accompanying drawings, in which:

FIG. 1 shows the relationship between the yield, on a hydrocarbon basis,of various C₂ to C₃ hydrocarbon constituents in the effluent and inlettemperature in some Examples and Comparative Examples; and

FIG. 2 shows the relationship between the propylene/ethylene ratio inthe effluent and inlet temperature in some Examples and ComparativeExamples.

In accordance with the present invention, catalytic conversion of afeedstock containing at least one C₁ to C₆ aliphatic hetero compoundselected from alcohols, ethers, carbonyl compounds and mixture thereof,into an effluent containing light olefins, in particular ethylene andpropylene, and selectively into propylene.

The C₁ to C₆ aliphatic alcohols may be monohydric and straight orbranched and may be selected from methanol, ethanol, propanol andbutanol. The ethers may be C₂ to C₄ ethers selected from dimethyl ether,diethyl ether or methyl ethyl ether. The carbonyl compounds may be C₂ toC₄ carbonyl compounds selected from formaldehyde, dimethyl ketone, oracetic acid. The feedstock is most preferably selected from methanol,ethanol, dimethyl ether, diethyl ether and mixtures thereof, withmethanol being particularly preferred.

In accordance with the process of the invention, the hydrocarbonfeedstocks are selectively converted in the presence of an MFI-type orMEL-type catalyst so as to produce propylene in the resultant effluent.The catalyst and process conditions are selected whereby the process hasa particular yield towards propylene in the effluent.

In accordance with a preferred aspect of the present invention, thecatalyst comprises a crystalline silicate of the MFI family which may bea zeolite, a silicalite or any other silicate in that family or the MELfamily which may be a zeolite or any other silicate in that family. Thethree-letter designations “MFI” and “MEL” each represent a particularcrystalline silicate structure type as established by the StructureCommission of the International Zeolite Association. Examples of MFIsilicates are ZSM-5 and silicalite. An example of an MEL zeolite isZSM-11 which is known in the art. Other examples are Boralite D, andsilicalite-2 as described by the International Zeolite Association(Atlas of zeolite structure types, 1987, Butterworths).

The preferred crystalline silicates have pores or channels defined byten oxygen rings and a high silicon/aluminium atomic ratio.

Crystalline silicates are microporous crystalline inorganic polymersbased on a framework of XO₄ tetrahedra linked to each other by sharingof oxygen ions, where X may be trivalent (e.g. Al, B, . . . ) ortetravalent (e.g. Ge, Si, . . . ). The crystal structure of acrystalline silicate is defined by the specific order in which a networkof tetrahedral units are linked together. The size of the crystallinesilicate pore openings is determined by the number of tetrahedral units,or, alternatively, oxygen atoms, required to form the pores and thenature of the cations that are present in the pores. They possess aunique combination of the following properties: high internal surfacearea; uniform pores with one or more discrete sizes; ionexchangeability; good thermal stability; and ability to adsorb organiccompounds. Since the pores of these crystalline silicates are similar insize to many organic molecules of practical interest, they control theingress and egress of reactants and products, resulting in particularselectivity in catalytic reactions. Crystalline silicates with the MFIstructure possess a bi-directional intersecting pore system with thefollowing pore diameters: a straight channel along [010]: 0.53-0.56 nmand a sinusoidal channel along [100]: 0.51-0.55 nm. Crystallinesilicates with the MEL structure possess a bi-directional intersectingstraight pore system with straight channels along [100] having porediameters of 0.53-0.54 nm.

The crystalline silicate catalyst has structural and chemical propertiesand is employed under particular reaction conditions whereby thecatalytic conversion to form light olefins, in particular propylene,readily proceeds.

The catalyst has a high silicon/aluminium atomic ratio, whereby thecatalyst has relatively low acidity. In this specification, the term“silicon/aluminium atomic ratio” is intended to mean the Si/Al atomicratio of the overall material, which may be determined by chemicalanalysis. In particular, for crystalline silicate materials, the statedSi/Al ratios apply not just to the Si/Al framework of the crystallinesilicate but rather to the whole material.

Different reaction pathways can occur on the catalyst. Hydrogen transferreactions are directly related to the strength and density of the acidsites on the catalyst, and such reactions arm preferably suppressed bythe use of high Si/Al ratios so as to avoid the formation of coke duringthe conversion process, thereby increasing the stability of thecatalyst. Moreover, the use of high Si/Al atomic ratios has been foundto increase the propylene selectivity of the catalyst, i.e. to reducethe amount of propane produced and/or to increase the propylene/ethyleneratio. This increases the purity of the resultant propylene.

In accordance with one aspect, a first type of MFI catalyst has a highsilicon/aluminum atomic ratio of from 250 to 500, whereby the catalysthas relatively low acidity. Hydrogen transfer reactions are directlyrelated to the strength and density of the acid sites on the catalyst,and such reactions are preferably suppressed so as to avoid theprogressive formation of coke which in turn would otherwise decrease thestability of the catalyst over time. Such hydrogen transfer reactionstend to produce saturates such as paraffins, intermediate unstabledienes and cyclo-olefins, and aromatics, none of which favoursconversion into light olefins. Cyclo-olefins are precursors of aromaticsand coke-like molecules, especially in the presence of solid acids, i.e.an acidic solid catalyst. The acidity of the catalyst can be determinedby the amount of residual ammonia on the catalyst following contact ofthe catalyst with ammonia which adsorbs to the acid sites on thecatalyst with subsequent ammonium desorption at elevated temperaturemeasured by differential thermogravimetric analysis.

With such high silicon/aluminum ratio in the crystalline silicatecatalyst, a stable conversion of the hydrocarbon feedstock can beachieved, with a high propylene yield of from 20 to 90%, more preferablyfrom 30 to 50%. The propylene selectivity is such that in the effluentthe propylene/ethylene weight ratio is typically from 2 to 10 and/or thepropylene/propane weight ratio is typically from 97/3 to 99.9/0.1. Suchhigh silicon/aluminum ratios in the catalyst reduce the acidity of thecatalyst, thereby also increasing the stability of the catalyst.

The MFI catalyst having a high silicon/aluminum atomic ratio for use inthe catalytic conversion process of the present invention ismanufactured by removing aluminum from a commercially availablecrystalline silicate. A typical commercially available silicalite has asilicon/aluminum atomic ratio of around 120. The commercially availableMFI crystalline silicate is modified by a steaming process which reducesthe tetrahedral aluminum in the crystalline silicate framework andconverts the aluminum atoms into octahedral aluminum the form ofamorphous alumina. Although in the steaming step aluminum atoms arechemically removed from the crystalline silicate framework structure toform alumina particles, those particles cause partial obstruction of thepores or channels in the framework. This inhibits the conversionprocesses of the present invention. Accordingly, following the steamingstep, the crystalline silicate is subjected to an extraction stepwherein amorphous alumina is removed from the pores and the microporevolume is, at least partially, recovered. The physical removal, by aleaching step, of the amorphous alumina from the pores by the formationof a water-soluble aluminum complex yields the overall effect ofde-alumination of the MFI crystalline silicate. In this way by removingaluminum from the MFI crystalline silicate framework and then removingalumina formed therefrom from the pores, the process aims at achieving asubstantially homogeneous de-alumination throughout the whole poresurfaces of the catalyst. This reduces the acidity of the catalyst, andthereby reduces the occurrence of hydrogen transfer reactions in theconversion process. The reduction of acidity ideally occurssubstantially homogeneously throughout the pores defined in thecrystalline silicate framework. This is because in the hydrocarbonconversion process hydrocarbon species can enter deeply into the pores.Accordingly, the reduction of acidity and thus the reduction in hydrogentransfer reactions which would reduce the stability of the MFI catalystare pursued throughout the whole pore structure in the framework. Theframework silicon/aluminum ratio is increased by this process to a valueof from 250 to 500.

Instead of an MFI-type catalyst, the process of the invention may use anMEL-type crystalline silicate having a silicon/aluminium atomic ratio offrom 150 to 800 which has been subjected to a steaming step. Inaccordance with this further aspect, an MEL catalyst for use in thecatalytic hydrocarbon conversion process may be manufactured by steamingan as-synthesised or commercially available crystalline silicate. TheMEL crystalline silicate catalyst for use in the invention mosttypically comprises a ZSM-11 catalyst which may be synthesised eitherusing diaminooctane as the templating agent and sodium silicate as thesilicon source or tetrabutyl phosphonium bromide as the templating agentand a silica sol as the silicon source. Thus the ZSM-11 catalyst may beprepared by mixing sodium silicate with 1,8 diaminooctane together withaluminium sulphate to form a hydrogel which is then allowed tocrystallise to form the crystalline silicate. The organic templatematerial is then removed by calcining. Alternatively, the ZSM-11catalyst is produced by reacting tetrabutyl phosphonium bromide andsodium hydroxide together with the silica sol prepared from colloidalsilica. Again, a crystallisation is performed to produce the crystallinesilicate and then the product is calcined.

In order to reduce the sodium content of the MEL crystalline silicate,the crystalline silicate is subjected to an ion exchange with a salt.Thereafter the material is dried. Typically, the crystalline silicate issubjected to ion exchange with ammonium ions, for example by immersingthe crystalline silicate in an aqueous solution of NH₄Cl or NH₄NO₃. Suchan ion exchange step is desirable if the amount of sodium ions presentin the crystalline silicate is so high that a crystalline sodiumsilicate phase is formed following calcination of the crystallinesilicate which would be difficult to remove.

-   -   The initial MEL crystalline silicate is modified by a steaming        process which, without being bound by theory, is believed to        reduce the tetrahedral aluminium in the crystalline silicate        framework and to convert the aluminium atoms into octahedral        aluminium in the form of amorphous alumina. Although in the        steaming step aluminium atoms are chemically removed from the        MEL crystalline silicate framework structure to form alumina        particles, those particles appear not to migrate and so do not        cause partial obstruction of the pores or channels in the        framework which would otherwise inhibit the conversion processes        of the present invention. The steaming step has been found to        improve significantly the propylene yield, propylene selectivity        and catalyst stability in the catalytic conversion process.

The steam treatment on the MEL catalyst is conducted at elevatedtemperature, preferably in the range of from 425 to 870° C., morepreferably in the range of from 540 to 815° C. and at atmosphericpressure and at a water partial pressure of from 13 to 200 kPa.Preferably, the steam treatment is conducted in an atmosphere comprisingfrom 5 to 100% steam. The steam treatment is preferably carried out fora period of from 1 to 200 hours, more preferably from 20 hours to 100hours. As stated above, the steam treatment tends to reduce the amountof tetrahedral aluminium in the crystalline silicate framework, byforming alumina.

Following the steaming step, the MEL catalyst is thereafter calcined,for example at a temperature of from 400 to 800° C. at atmosphericpressure for a period of from 1 to 10 hours.

Following the steaming step, the MEL catalyst may be contacted by acomplexing agent for aluminium which may comprise an acid in an aqueoussolution thereof or a salt of such an acid or a mixture of two or moresuch acids or salts. The complexing agent may in particular comprise anamine, such as ethyl diamine tetraacetic acid (EDTA) or a salt thereof,in particular the sodium salt thereof. Following the contacting of theMEL crystalline silicate by the complexing agent, the crystallinesilicate may be subjected to a second ion exchange step for reducing thesodium content of the crystalline silicate still further, for example bycontacting the catalyst with an ammonium nitrate solution.

The MEL or MFI crystalline silicate catalyst may be mixed with a binder,preferably an inorganic binder, and shaped to a desired shape, e.g.extruded pellets. The binder is selected so as to be resistant to thetemperature and other conditions employed in the catalyst manufacturingprocess and in the subsequent catalytic conversion process. The binderis an inorganic material selected from clays, silica, metal oxides suchas Zr0₂ and/or metals, or gels including mixtures of silica and metaloxides. The binder is preferably alumina-free. However, aluminium incertain chemical compounds as in AlPO₄'s may be used as the latter arequite inert and not acidic in nature. If the binder which is used inconjunction with the crystalline silicate is itself catalyticallyactive, this may alter the conversion and/or the selectivity of thecatalyst. Inactive materials for the binder may suitably serve asdiluents to control the amount of conversion so that products can beobtained economically and orderly without employing other means forcontrolling the reaction rate. It is desirable to provide a catalysthaving a good crush strength. This is because in commercial use, it isdesirable to prevent the catalyst from breaking down into powder-likematerials. Such clay or oxide binders have been employed normally onlyfor the purpose of improving the crush strength of the catalyst. Aparticularly preferred binder for the catalyst of the present inventioncomprises silica.

The relative proportions of the finely divided crystalline silicatematerial and the inorganic oxide matrix of the binder can vary widely.Typically, the binder content ranges from 5 to 95% by weight, moretypically from 20 to 50% by weight, based on the weight of the compositecatalyst. Such a mixture of crystalline silicate and an inorganic oxidebinder is referred to as a formulated crystalline silicate.

In mixing the catalyst with a binder, the catalyst may be formulatedinto pellets, extruded into other shapes, or formed into a spray-driedpowder.

Typically, the binder and the crystalline silicate catalyst are mixedtogether by an extrusion process. In such a process, the binder, forexample silica, in the form of a gel is mixed with the crystallinesilicate catalyst material and the resultant mixture is extruded intothe desired shape, for example pellets. Thereafter, the formulatedcrystalline silicate is calcined in air or an inert gas, typically at atemperature of from 200 to 900° C. for a period of from 1 to 48 hours.

The binder preferably does not contain any aluminium compounds, such asalumina. This is because as mentioned above the preferred catalyst has aselected silicon/aluminium ratio of the crystalline silicate. Thepresence of alumina in the binder yields other excess alumina if thebinding step is performed prior to the aluminium extraction step. If thealuminium-containing binder is mixed with the crystalline silicatecatalyst following aluminium extraction, this re-aluminates thecatalyst. The presence of aluminium in the binder would tend to reducethe propylene selectivity of the catalyst, and to reduce the stabilityof the catalyst over time.

In addition, the mixing of the catalyst with the binder may be carriedout either before or after any steaming step.

The various preferred catalysts have been found to exhibit highstability, in particular being capable of giving a stable propyleneyield over several days, e.g. up to ten days. This enables the catalyticconversion process to be performed continuously in two parallel “swing”reactors wherein when one reactor is operating, the other reactor isundergoing catalyst regeneration. The catalyst also can be regeneratedseveral times. The catalyst is also flexible in that it can be employedto crack a variety of feedstocks, either pure or mixtures, coming fromdifferent sources in the oil refinery or petrochemical plant and havingdifferent compositions.

In the catalytic conversion process, the process conditions are selectedin order to provide high selectivity towards propylene, a stableconversion into propylene over time, and a stable product distributionin the effluent. Such objectives are favoured by the use of a low aciddensity in the catalyst (i.e. a high Si/Al atomic ratio) in conjunctionwith a low pressure, a high inlet temperature and a short contact time,all of which process parameters are interrelated and provide an overallcumulative effect (e.g. a higher pressure may be offset or compensatedby a yet higher inlet temperature). The process conditions are selectedto disfavour hydrogen transfer reactions leading to the formation ofparaffins, aromatics and coke precursors. The process operatingconditions thus employ a high space velocity, a low pressure and a highreaction temperature.

The weight hourly space velocity (WHSV) with respect to theoxygen-containing hydrocarbon feedstock ranges from 0.5 to 30 h⁻¹,preferably from 1.0 to 20 h⁻¹. The oxygen-containing hydrocarbonfeedstock is preferably fed at a total inlet pressure sufficient toconvey the feedstock through the reactor. Preferably, the total absolutepressure in the reactor ranges from 0.5 to 10 bars. The oxygenatedpartial pressure ranges from 20 to 400 kPa, preferably from 50 to 200kPa. A particularly preferred oxygenated partial pressure is 100 kPa.The oxygenates feedstocks may be fed undiluted or diluted with steam,e.g. from 0 to 80 wt % steam, typically about 30 wt % steam, and/or inan inert gas, e.g. nitrogen or hydrogen. The use of a low oxygenatespartial pressure, for example atmospheric pressure, tends to lower theincidence of hydrogen transfer reactions in the conversion process,which in turn reduces the potential for coke formation which tends toreduce catalyst stability. Preferably, the inlet temperature of thefeedstock ranges from 350 to 650° C., more preferably from 400 to 600°C., yet more preferably from 450 to 585° C., typically around 450° C. to550° C.

The catalytic conversion process can be performed in a fixed bedreactor, a moving bed reactor or a fluidized bed reactor. A typicalfluid bed reactor is one of the FCC type used for fluidized-bedcatalytic cracking in the oil refinery. A typical moving bed reactor isof the continuous catalytic reforming type. As described above, theprocess may be performed continuously using a pair of parallel “swing”fixed bed reactors.

Since the catalyst exhibits high stability for an extended period,typically at least around ten days, the frequency of regeneration of thecatalyst is low. More particularly, the catalyst may accordingly have alifetime which exceeds one year.

The light fractions of the effluent, namely the C₂ and C₃ cuts, cancontain more than 90% olefins (i.e. ethylene and propylene). Such cutsare sufficiently pure to constitute chemical grade olefin feedstocks.The propylene yield in such a process can range from 20 to 90%. Thepropylene/ethylene weight ratio typically ranges from 2 to 10, moretypically from 2 to 5. The propylene/propane weight ratio typicallyranges from 10 to 1000, more typically from 15 to 100. These two ratiosmay be higher than obtainable using prior art processes describedherein. The propylene/aromatics weight ratio may range from 2.5 to 100,more typically from 3 to 10.

In accordance with the present invention therefore, hydrocarbonfeedstocks containing at least one C₁ to C₆ aliphatic hetero compoundselected from alcohols, ethers, carbonyl compounds and mixtures thereofare subject to a catalytic conversion process which selectively formspropylene as well as ethylene, and thereafter, the effluent is separatedinto a C₂ and C₃ combined product that is recovered in a commonfractionation train, and into a C₄+ product. The C₂ and C₃ combinedproduct is high in propylene, and relatively low in ethylene andpropane.

The present invention will now be described in greater detail withreference to the following non-limiting Examples.

EXAMPLE 1

In Example 1, a laboratory scale fixed bed reactor had provided thereina crystalline silicate catalyst of the MFI-type. The catalyst comprisessilicalite which had a silicon/aluminium atomic ratio of 273 and hadbeen produced by a de-alumination process as described hereinabove.

More specifically, the silicalite catalyst was prepared by steaming 4.2kg of silicalite at 550° C. for a period of 48 hours with steam in arotating laboratory furnace. Thereafter, 2 kg of the steamed silicalitewas then treated with an aqueous solution of the sodium salt of ethyldiamine tetraacetic acid (EDTA-Na₂), there being 8.4 litres of a 0.055molar solution thereof for the 2 kg of silicalite. The treatment was fora period of 18 hours at boiling temperature. The silicalite was thensubsequently filtered and washed thoroughly with de-ionised water. Thisprocess extracted aluminium from the silicalite.

Thereafter, an extruded catalyst was prepared using a kneader, inparticular a Guittard type M5 No. 2295 kneader. In particular, 1640 g ofthe treated silicalite, 112 g of silica powder (Degussa FK500) and 726 gof silica sol (Nyacol 2040 from EKA containing about 41% silica byweight) were mixed for a few minutes to homogonize them, and then 600 mlof distilled water was added to the mixture to obtain a paste, which wasthen mixed for another 30 minutes. After the 30 minute mixing time, 10 gof polyelectrolyte solution (Nalco 9779) were added to the mixture andkneaded for 1 minute. Then 30 g of methyl-hydroxy-ethyl-cellulose(Tylose from Hoechst MHB1000P2) were added. The loss on ignition (LOI)was about 33 wt %. The extruder (Alexanderwerk type AGMR No. 04231162)was equipped with a die plate aperture of 2.5 mm, which was quadralobeshaped. The paste was passed 2 to 3 times through the extruder. Theresultant extrudates were air-dried over night, then dried at 110° C.for 16 hours in a drying oven with a heating rate of 60° C. per hour,and then calcined at a temperature of 600° C. for a period of 10 hours.Finally, the catalyst was subjected to ion-exchange, whereby 1740 g ofthe extruded catalyst was ion-exchanged using NH₄Cl (0.5 molar and 7310ml of solution) twice, the first time being for a period of 18 hours andthe second time being for a period of 3 hours, both at the boilingtemperature of the solution. Finally, the catalyst was filtered off,washed and calcined at a temperature of 400° C. for a period of 3 hours.

The resultant modified silicalite catalyst was in the form of particlesof crushed extrudates of 35 to 45 mesh size. Chemical analysis of thecatalyst indicate that the composition as SiO₂ 99.594 wt %, Al₂O₃ 0.310wt %, Na₂O 0.028 wt % and Fe₂O₃ 0.058 wt %. This provided asilicon/aluminium atomic ratio of 273.

The laboratory scale reactor had a diameter of 10 mm and was loaded witha catalyst load of 3 g. The reactor was subjected to a pre-treatment at500° C. under nitrogen gas overnight. The reactor was operated atatmospheric pressure. The reactor was fed with an oxygenates feedstockcomprising 70 wt % methanol and 30 wt % water, in the form of steam, ata methanol partial pressure of 56 kPa. The WHSV, with respect to themethanol, was 1.9 h⁻¹. The total time on stream [TOS] was 457 minutes.Initially, the reactor inlet temperature was −400° C. and after 270minutes on stream, the reactor inlet temperature was increased to 450°C. The composition of the effluent is shown in Table 1. The compositionof the effluent was analysed using an on-line apolar column (DB-1, 0.4micron, JNW ScientificCat. No. 1271043).

It may be seen from Table 1 that the methanol was 100% convertedthroughout the time on stream. At a reactor inlet temperature of around400° C., the propylene yield was around 22 wt % and the ethylene yieldwas around 11%. The propane yield was around 1.4 wt %. When the reactortemperature was increased to 450° C., the propylene yield was increasedto around 30 wt %, the ethylene yield decreased slightly to less than 10wt %, and the propane yield decreased slightly as well. At 450° C., thepropylene/ethylene weight ratio was about 3 or greater and thepropylene/propane weight ratio was about 23 or greater. Accordingly, inthis Example, the propylene selectivity was high, and the relativelyhigh values of the propane/ethylene weight ratio and thepropylene/propane weight ratio provided high propylene purity in afractionated C₂ and C₃ combined cut.

The yields, on a hydrocarbon basis (water free), of propylene, ethyleneand propane at the two temperatures in Example 1 are shown in FIG. 1.The propylene/ethylene weight ratios at the two temperatures in Example1 are shown in FIG. 2.

EXAMPLE 2

In Example 2 the process of Example 1 was repeated with the samefeedstock, catalyst and WHSV but at a higher reactor inlet temperatureof 550° C. The results are shown in Table 2.

The Example was carried out for a total of 185 minutes on stream at atemperature of 550° C.

It may be seen that the propylene yield is increased at the highertemperature of 550° C. as compared to the temperatures of Example 1. Thepropylene yield was about 40 wt % after 185 minutes on stream. At thattime, the propylene/ethylene weight ratio was about 3.3 and thepropylene/propane weight ratio was about 38. Again, this indicates notonly high propylene selectivity, but high propylene purity in afractionated C₂ and C₃ combined cut. Like Example 2, Example 1 showshigh stability of the catalyst when used in a fixed bed reactor overtime.

The yields, on a hydrocarbon basis (water free), of propylene, ethyleneand propane at the temperature in Example 2 are shown in FIG. 1. Thepropylene/ethylene weight ratios at the temperature in Example 2 areshown in FIG. 2.

COMPARATIVE EXAMPLE 1

In this Comparative Example, Example 1 was repeated using a differentcatalyst, namely a silica-alumina-phosphate catalyst, in particular SAPO34 available from UOP of Des Plaines, Ill., USA, having a particle sizeof 35-45 mesh. The same feedstock and WHSV were employed as in Examples1 and 2. The reactor temperature was a constant 450° C. A maximum timeon stream was 211 minutes. The results are shown in Table 3.

As may be seen from Table 3, initially the propylene yield was higherthan the ethylene yield but the propylene/ethylene weight ratio rapidlydecreased below unity. Therefore the propylene selectivity of thiscatalyst is lower than that employed in the present invention. Moreover,after only 149 minutes on stream the methanol conversion fell below 100%and the effluent included the methanol from the feedstock as well asdimethyl ether. This shows that the SAPO 34 catalyst when used in afixed bed had a low stability.

The propylene/ethylene weight ratios for the catalyst of ComparativeExample 1 are shown in FIG. 2.

COMPARATIVE EXAMPLE 2

In this Comparative Example, Comparative Example 1 was repeated using adifferent catalyst, the catalyst being a silicalite available incommerce under product number S-115 Na-6 from UOP of Des Plaines, Ill.,USA, the silicalite having a silicon/aluminium atomic ratio of 177. Thesilicalite had a chemical composition of SiO₂ 99.450 wt %, Al₂O₃ 0.478wt %, Na₂O 0.006 wt % and Fe₂O₃ 0.052 wt %, yielding a silicon/aluminiumratio of 177. The silicalite was in the form of particles of 35 to 45mesh. The WHSV was 109 h⁻¹ as in Examples 1 and 2 and in ComparativeExample 1 and the feed also comprised 70 wt % methanol and 30 wt %steam. The process of Comparative Example 2 was carried out at tworeactor inlet temperatures, namely at a temperature of 450° C. for up to208 minutes on stream, and at a temperature of 500° C. thereafter up toa total time on stream of 380 minutes. The results are summarised inTable 4.

From Table 4, it may be seen that while the stability of the catalyst ishigher as compared to Comparative Example 1, the propylene selectivityand purity are less than obtained in accordance with Examples 1 and 2.Thus at the same comparison temperature of 450° C., in ComparativeExample 2 the propylene yield was consistently less than 30 wt %, lowerthan that achievable in Example 1 at the corresponding temperature.Moreover, at that temperature of 450° C., the propylene/ethylene weightratio was about 2.7, lower than obtainable in Example 1. Furthermore, inComparative Example 2, the propylene/propane weight ratio at a reactorinlet temperature of 450° C. was about 9 or less, thereby indicatinglower propylene purity corresponding to that obtainable using thecorresponding temperature in Example 1. In Comparative Example 2 in yeta higher reactor inlet temperature of 500° C., the ethylene yield andthe propane yield were higher than that obtainable in Example 1 of 450°C.

The yields, on a hydrocarbon basis (water free), of propylene, ethyleneand propane at the two temperatures in Comparative Example 2 are shownin FIG. 1. The propylene/ethylene weight ratios at the two temperaturesin Comparative Example 2 are shown in FIG. 2.

COMPARATIVE EXAMPLE 3

In this Comparative Example, Comparative Example 1 was repeated butusing a feed comprising 100 wt % methanol. The same WHSV and reactortemperature were employed as in Comparative Example 1. The maximum timeon stream was 102 minutes. The results are shown in Table 5.

Table 5 shows that for Comparative Example 3, although the propyleneyield is more stabilised compared to Comparative Example 1, thepropylene/ethylene ratio rapidly decreased below unity, and therefore islower than that achievable using Examples 1 and 2.

TABLE 1 Example 1 TOS [min] 145 270 332 395 457 Temperature [° C.] 400400 450 450 450 Conversion [%] 100 100 100 100 100 Yields [wt %] C1 1.441.11 4.76 2.96 3.30 C2− 11.20 10.96 10.19 9.88 9.44 C2 0.12 0.12 0.140.13 0.11 C3− 22.06 22.52 29.25 30.61 30.95 C3 1.49 1.38 1.51 1.30 1.34C4's 4.71 4.19 3.05 2.65 2.76 C4−'s 19.06 16.05 18.80 18.85 19.24 C5+'s39.92 43.66 32.30 33.62 32.85 Total 100.00 100.00 100.00 100.00 100.00Aromatics C6 0.49 0.49 0.81 0.82 0.79 C7 0.99 0.90 1.32 1.36 1.11 C86.13 5.69 5.61 6.03 5.12 Total aromatics 7.61 7.08 7.74 8.21 7.03

TABLE 2 Example 2 TOS [min] 61 123 185 Temperature [° C.] 550 550 550Conversion [%] 100 100 100 Yields [wt %] C1 7.78 4.99 4.68 C2− 12.6812.69 12.30 C2 0.30 0.31 0.30 C3− 36.25 39.81 40.50 C3 0.95 1.08 1.07C4's 0.90 0.74 0.69 C4−'s 13.60 15.56 16.12 C5+'s 27.55 24.83 24.34Total 100.00 100.00 100.00 Aromatics C6 0.44 0.51 0.54 C7 3.47 3.45 3.20C8 7.59 6.87 6.96 Total aromatics 11.50 10.84 10.70

TABLE 3 Comparative Example 1 TOS [min] 24 86 149 211 Temperature [° C.]450 450 450 450 Conversion [%] 100 100 93.6 78.6 Yields [wt %] C1 2.094.29 5.22 0 C2− 32.83 42.86 33.02 0 C2 0.78 0.86 0.69 0 C3− 35.78 37.2127.84 0 C3 3.26 0.97 0.57 0 DME 0 0 15.38 77.92 Methanol 0 0 6.37 21.36C4's 0.43 0.08 0.13 0 C4−'s 15.93 10.11 7.35 0.72 C5+'s 8.84 3.55 2.28 0Total 99.95 99.93 98.84 100.00

TABLE 4 Comparative Example 2 TOS [min] 20 137 199 208 324 386Temperature [° C.] 450 450 450 450 500 500 Conversion [%] 100 100 100100 100 100 Yields [wt %] C1 3.03 2.29 1.69 1.77 4.40 4.33 C2− 9.71 9.639.58 10.21 14.27 14.59 C2 0.30 0.32 0.32 0.32 0.60 0.63 C3− 28.14 25.6626.02 27.87 30.89 31.09 C3 3.13 3.70 3.42 3.07 2.95 2.98 C4's 4.48 5.504.96 4.16 2.39 2.35 C4−'s 21.86 21.22 21.18 21.48 18.10 17.67 C5+'s29.31 31.61 32.80 31.06 26.33 26.30 Total 100.0 100.0 100.0 100.0 99.999.9 Aromatics C6 2.12 2.29 2.36 2.95 1.11 1.06 C7 2.44 2.51 3.02 2.864.26 4.39 C8 5.82 5.82 6.73 7.38 9.15 9.52 Total aromatics 10.38 10.6212.12 13.20 14.52 14.98

TABLE 5 Comparative Example 3 TOS [min] 20 61 102 Temperature [° C.] 450450 450 Conversion [%] 100 100 100 Yields [wt %] C1 13.09 8.19 8.73 C2−30.12 41.33 42.64 C2 0.56 0.74 0.77 C3− 33.03 34.69 34.76 C3 2.36 1.150.78 DME 0.00 0.00 0.00 Methanol 0.00 0.00 0.00 C4's 0.37 0.10 0.06 C4−'s 13.59 10.20 8.91 C5+'s 6.81 3.54 3.20 Total 99.93 99.93 99.85 C3−/C2−1.10 0.84 0.82

1-10. (canceled)
 11. A process for converting a carbon containingfeedstock to provide an effluent containing light olefins, the processcomprising: passing a carbon containing feedstock through a reactorcontaining a crystalline silicate catalyst to produce an effluentcomprising propylene which is recovered from the reactor, wherein thecarbon containing feedstock is passed over the crystalline silicatecatalyst at a reactor inlet temperature of about 550° C.; wherein thecarbon containing feedstock contains at least one C₁ to C₆ aliphatichetero compound selected from a group consisting of alcohols, ethers,carbonyl compounds and mixtures thereof, and steam in an amount up to 80weight % of the carbon containing feedstock; and wherein the crystallinesilicate catalyst is pretreated by subjecting the crystalline silicatecatalyst to de-alumination by steaming, and wherein the crystallinesilicate catalyst is selected from at least one of an MFI-typecrystalline silicate having a silicon/aluminum atomic ratio of from 250to 500 and an MEL-type crystalline silicate having a silicon/aluminumatomic ratio of from 150 to
 800. 12. The process of claim 11, whereinthe carbon containing feedstock contains at least one hetero compoundselected from the group consisting of methanol, ethanol, dimethyl ether,diethyl ether and mixtures thereof.
 13. The process of claim 11, whereinthe carbon containing feedstock is passed over the crystalline silicatecatalyst at a WHSV of from 0.5 to 30 h⁻¹, the WHSV being based on theweight of the at least one C₁ to C₆ aliphatic hetero compound in thefeedstock.
 14. The process of claim 11, wherein the partial pressure ofthe at least on C₁ to C₆ aliphatic hetero compound, when passed over thecrystalline silicate catalyst, is from 20 to 400 kPa.
 15. The process ofclaim 11, wherein the effluent exhibits a propylene/ethylene weightratio ranging from 2 to
 10. 16. The process of claim 11, wherein theeffluent exhibits a propylene/ethylene weight ratio ranging from 2 to 5.17. The process of claim 11, wherein the propylene yield is above 35%.18. The process of claim 17, wherein the time on stream is about 61minutes.
 19. The process of claim 17, wherein the time on stream isabout 123 minutes.
 20. The process of claim 17, wherein the time onstream is about 185 minutes.